Fluidized bed biogasifier and method for gasifying biosolids

ABSTRACT

A fluidized bed biogasifier is provided for gasifying biosolids. The biogasifier includes a reactor vessel and a feeder for feeding biosolids into the reactor vessel at a desired feed rate during steady-state operation of the biogasifier. A fluidized bed in the base of the reactor vessel has a cross-sectional area that is proportional to at least the fuel feed rate such that the superficial velocity of gas is in the range of 0.1 m/s (0.33 ft/s) to 3 m/s (9.84 ft/s). In a method for gasifying biosolids, biosolids are fed into a fluidized bed reactor. Oxidant gases are applied to the fluidized bed reactor to produce a superficial velocity of producer gas in the range of 0.1 m/s (0.33 ft/s) to 3 m/s (9.84 ft/s). The biosolids are heated inside the fluidized bed reactor to a temperature range between 900° F. (482.2° C.) and 1700° F. (926.7° C.) in an oxygen-starved environment having a sub-stoichiometric oxygen level, whereby the biosolids are gasified.

CROSS REFERENCE TO RELATED APPLICATIONS

This application is a divisional of application Ser. No. 13/361,582,filed Jan. 30, 2012, (which is hereby incorporated by reference).

This application includes material which is subject to copyrightprotection. The copyright owner has no objection to the facsimilereproduction by anyone of the patent disclosure, as it appears in thePatent and Trademark Office files or records, but otherwise reserves allcopyright rights whatsoever

FIELD

The present invention relates in general to the field of sewage sludgetreatment, and in particular to a fluidized bed biogasification systemand method for use in treatment of biosolids from sewage sludge.

BACKGROUND

U.S. Pat. No. 7,793,601 to Davison et al. entitled “Side Feed/Centre AshDump System” describes a gasifier for gasifying solid fuel, includingsolid organic materials. The gasifier includes a primary oxidationchamber having an inner surface lined with a refractory material. Aninlet opening in one of the sides is provided for infeed of the solidfuel into the primary oxidation chamber. A storage container stores thesolid fuel, and a transfer means connects the storage container with theinlet opening for transferring in an upwardly inclined direction thesolid fuel from the storage container through the inlet opening into theprimary oxidation chamber to form an upwardly mounted fuel bed of thesolid fuel including the organic materials on the bottom of the primaryoxidation chamber. The transfer means includes a hydraulic ram feederand a compression tube, the hydraulic ram feeder driving fuel from thestorage container into the compression tube, thereby compacting thefuel. Means are provided for supplying an oxidant into the primaryoxidation chamber to gasify the solid organic materials to produce agaseous effluent, thereby leaving a residue of solid fuel. Means areprovided for removing the gaseous effluent from the primary oxidationchamber. An opening in the bottom of the primary oxidation chamber hasmounted thereunder a means for the removal of the residue, including awalking-floor feeder.

U.S. Pat. No. 7,322,301 to Childs describes a method for processing wetsewage sludge or other feedstock including carbonaceous materialprincipally composed of wet organic materials in a gasifier to produceuseful products. The sludge or feedstock is first dewatered usingthermal energy in a location separate from the gasifier. The feedstockis processed with a small amount of oxygen or air present at atemperature required to break down the feedstock and generate producergas and char in the gasifier. Some of the fuel produced during thefeedstock processing step is fed back to the separate location andburned to provide the thermal energy required in the feedstockdewatering step and thereby minimize or eliminate the need for externalenergy to dry the wet feedstock.

U.S. Pat. No. 6,120,567 to Cordell et al. describes a method forgasifying solid organic materials to produce a gaseous effluent andsolid residue. A primary oxidation chamber having a converging upperportion and a bottom portion are provided. Solid organic materials areintroduced into the primary oxidation chamber upwardly from the bottomportion of the primary oxidation chamber to provide a mass of solidorganic materials in the primary oxidation chamber. The mass of solidorganic materials is heated in the primary oxidation chamber. An oxidantis added to the primary oxidation chamber to gasify the heated mass ofsolid organic materials in the primary oxidation chamber and to initiatea flow of gaseous effluent within the primary oxidation chamber. Agaseous effluent flow path is established within the primary oxidationchamber, whereby a portion of the gaseous effluent repeatedly flows in arecirculating upward and downward direction through the heated solidorganic materials to enhance continuous oxidation of the solid organicmaterials. A further portion of the gaseous effluent flow is advanced ina direction outward from the primary oxidation chamber. The solidresidue is then transferred out of the primary oxidation chamber.

SUMMARY

In an embodiment, a fluidized bed biogasifier is provided for gasifyingbiosolids. The biogasifier includes a reactor vessel and a feeder forfeeding biosolids into the reactor vessel at a desired feed rate duringsteady-state operation of the biogasifier. A fluidized media bed in thebase of the reactor vessel has a cross-sectional area that isproportional to at least the fuel feed rate so as to produce asuperficial velocity of gas in the range of 0.1 m/s (0.33 ft/s) to 3 m/s(9.84 ft/s). In an embodiment, the internal diameter of the reactor isconfigured to be small enough to ensure that the fluidized bed is ableto be fluidized adequately for the desired fuel feed rate and the flowrate of the fluidizing gas mixture at different operating temperatures,but not so small as to create such high gas velocities that a sluggingfluidization regime occurs and media is projected up the freeboardsection. Other factors may be used in the design and sizing of thebiogasifier, including internal diameter of the bed section, internaldiameter of a freeboard section, height of the freeboard section, beddepth and the bed section height.

In an embodiment, a method for gasifying biosolids is provided.Biosolids are fed into a fluidized bed reactor. Fluidizing gasconsisting of air, flue gas, pure oxygen or steam, or a combinationthereof, is introduced into the fluidized bed reactor to create avelocity range inside the freeboard section of the gasifier that is inthe range of 0.1 m/s (0.33 ft/s) to 3 m/s (9.84 ft/s). The biosolids areheated inside the fluidized bed reactor to a temperature range between900° F. (482.2° C.) and 1700° F. (926.7° C.) in an oxygen-starvedenvironment having sub-stoichiometric levels of oxygen, e.g., typicallyoxygen levels of less than 45% of stoichiometric.

BRIEF DESCRIPTION OF THE DRAWINGS

The foregoing and other objects, features, and advantages of theinvention will be apparent from the following more particulardescription of preferred embodiments as illustrated in the accompanyingdrawings, in which reference characters refer to the same partsthroughout the various views. The drawings are not necessarily to scale,emphasis instead being placed upon illustrating principles of theinvention.

FIGS. 1A and 1B show schematic block diagrams illustrating alternateembodiments of the flow of an overall treatment process for biosolids.

FIG. 2 shows a schematic side view illustrating a fluidized bedbiogasifier in accordance with an embodiment of the invention.

FIG. 3 shows a perspective view illustrating the gas distributor of thebiogasifier in accordance with an embodiment of the invention.

FIGS. 4A-4B show a schematic block diagram illustrating part of agasification process in accordance with the invention in an embodiment.

FIG. 5 shows a schematic side view illustrating a non-limiting exampleof biogasifier internal dimensions in accordance with the invention inan embodiment.

DETAILED DESCRIPTION

Reference will now be made in detail to the preferred embodiments of thepresent invention, examples of which are illustrated in the accompanyingdrawings.

With reference to FIG. 1A, an overall process for treatment of biosolidsis shown. The process begins with a dewatering unit 1 for removing watercontent from the wet sewage sludge. Dewatering may be achieved in thedewatering unit 1 utilizing integrated technology including, e.g., abelt filter press, plate-and-frame press, and/or centrifuge. Specificvolumes and blending of polymer are used in accordance with the specificapplication to help flocculate the materials and gain the most efficientdewatered percentage prior to the next step in the overall process.

The material can then be fed into storage 1 and storage 2, which may besized such that appropriately staged volumes of dewatered cake aremaintained within the pipeline to ensure a continuous process in thedryer. The storage may be controlled through load cells or levelsensors, and the moisture of the biosolids may be constantly measureden-route to the pre-drying/dryer operation to ensure optimum control andperformance output.

A pre-dryer unit 3 then takes the condensate and steam, typically at215° F. to 240° F., from the biosolids dryer and uses it to pre-heat thebiosolids prior to entry into the dryer. This pre-dryer feed loop 6minimizes the amount of energy subsequently required by the dryer anddemand on the biogasifier to the thermal fluid loop.

The biosolids are then fed to a biosolids dryer 4, which may consist ofa continuous feed, screw-type biosolids dryer. The biosolids dryer mayutilize thermal energy produced from the coupled gasification-combustionsystem as its primary source of energy for use in the drying process.The flue gas may then be passed through the heat exchanger. The energyfrom the heated flue gas can be conveyed through a gas-to-liquid heatexchanger into the thermal heating fluid. This thermal fluid may berecirculated through the chambers of the dryer “indirectly” conveyingenergy to the biosolids and thereby drying the product. The flue gastemperature into the heat exchanger may be controlled through the use ofinduction and dilution fans in order to maintain a consistent heatsource for the drying system.

An odor control section 5 utilizes multiple types of odor controltechnology to extract all pollutants that may become airborne andpotentially harmful outside of the process.

Dried biosolids storage 7 can be provided between the biosolids dryer 4and the biogasifier feed system 8. The dried biosolids storage 7 may besized to ensure that appropriately staged volumes of fuel can bemaintained. This ensures a continuous process in the biogasifier. Thestorage 7 may be controlled through load cells or level sensors toensure optimum control and performance output.

The biogasifier feed system 8 can be configured to meter the delivery ofbiosolids fuel into the biogasifier. In an embodiment, the feed system 8includes two metering screws and a high-speed injection screw into thebiogasifier bed for this purpose.

The biogasifier feed system 8 feeds a biogasifier unit 9. In anembodiment, the biogasifier unit 9 is of the bubbling fluidized bed typewith a custom fluidizing gas delivery system and multiple instrumentcontrol. The biogasifier unit 9 provides the ability to continuouslyoperate, discharge ash and recycle flue gas for optimum operation. Thebiogasifier unit 9 can be designed to provide optimum control oftemperature, reaction rate and conversion of the biosolids fuel intoproducer gas. The control mechanism for temperature within thebiogasifier may be based on the system's ability to adjust the oxygencontent relative to the fuel feed rate, thus adjusting reactiontemperatures. In an embodiment, the biogasifier operates within atemperature range of 900°-1700° F. during steady state operation. Inanother embodiment, the biogasifier operates within a temperature rangeof 1150°-1600° F. during steady state operation, and this range issignificantly preferred since it has been determined that below 1150° F.there will be limited reactions occurring and above 1600° F. there willbe very pronounced agglomeration problems with biosolids. Oxygen is areactant that facilitates the chemical reactions necessary forsustaining the gasification process.

A cyclone separator 10 can be provided to separate material exhaustedfrom the fluidized bed reactor to create clean producer gas and ash fordisposal. In an embodiment, the cyclone separator 10 is efficient toensure over 95% particulate removal and gas clean up.

Ash discharge 11 from both the biogasifier 9 and the cyclone separator10 are safe to transport. Such ash discharge may be disposed of, or mayprovide value in commercial applications, such as from the recovery anduse of phosphorous.

A staged-combustion thermal oxidizer 12 may be provided for thermaloxidation of clean producer gas from the cyclone separator 10. Thisprocess may be used to generate heat that is used as energy in thesystem to operate, e.g., the biosolids dryer 4 in part or in full. Thethermal oxidizer may be a refractory lined steel unit with ports for theintroduction of air to promote the homogenous blending of the producergas with air, taking the resultant mass to combustion. The producergases are delivered from the biogasifier 9, to the thermal oxidizer 12where they are reacted in a multi-stage process under negative pressure.In each stage, the temperature and oxygen content are tightlycontrolled, resulting in conversion of the producer gas to a very clean,low NOx, high-grade flue gas stream that may be recovered and used, asnoted above, for the generation of heat as an energy input back into thesystem, or may be output from the system for other commercial uses.

Emissions controls 13 are provided for controlling emissions from thestaged combustion thermal oxidizer 12. The emissions controls 13 mayinclude an injection/misting point in which emission control chemicalsare inserted after the combustion stage to reduce the NOx levels withinthe exiting air stream. Emission control chemicals may be either one ora combination of aqueous ammonia and urea.

A bypass stack 14 may be utilized for exhausting hot flue gas and/orthermal energy in the case of an emergency upset within the system, ascontrolled via a Supervisory Control And Data Acquisition (SCADA)system.

Dependent upon the end use of the energy, the media into which thethermal energy is conveyed may be thermal oil, hot water, steam, or thelike. This process is achieved through the use of high efficiency heatexchangers that transfer thermal energy from the hot flue gas intoanother media. This converted media can be used to provide the energy tothe selected recovery system. The gasification-combustion processfollowed by energy recovery utilizes energy in biosolids and reduces oreliminates the need for use of fossil fuels to dry the biosolids. Thismakes the disclosed gasification system an energy-efficient andenvironmentally friendly solution to biosolids management.

A first heat exchanger 15 receives heated flue gas from the stagedcombustion thermal oxidizer 12. The first heat exchanger 15 may be agas-to-liquid type heat exchanger. Energy from the heated flue gas maybe conveyed through the first heat exchanger 15 into a thermal heatingfluid. This thermal fluid may be recirculated through the chambers ofthe dryer, indirectly conveying thermal energy to the biosolids andthereby drying the product. A thermal fluid loop 16 from the first heatexchanger 15 to the biosolids dryer 4 allows the energy recovery systemto utilize heated flue gas from the thermal oxidizer as its primarysource of energy.

A second heat exchanger 17 utilizes additional energy recovery streams(typically wasted) and returns the energy to the process for optimumperformance efficiency. The second heat exchanger 17 in this embodimentis dedicated to a recycle flue gas loop 18. The source of the oxygen forthe disclosed gasification system may be designed to come from air orflue gas, or mixtures thereof. In an embodiment, the disclosed system isconfigured such that oxygen may be introduced via re-circulated fluegas, which has an oxygen content of approximately 50% of ambient air.Having both ambient air and flue gas available as an oxygen sourceprovides a level of flexibility for oxygen delivery that can be furtherused to control the biogasifier 9 temperature more precisely.

The infusion of flue gas into the fluidizing air stream provides some ofthe gas velocity required to transport particles out of the biogasifier9 and into the cyclone, especially during low fuel feed rate operationconditions.

A third heat exchanger 19 may be provided to take additional energystreams, which are typically wasted in other processes, and re-insertthem into the process for greater performance efficiency. In anembodiment, heat exchanger 19 is dedicated to the hot air feed 20 intothe thermal oxidizer. The hot air feed 20, which comprises oxidizer airprovides heating of the air being injected (via staged combustion airrings) into the oxidizer which assists in the overall efficiency andcombustion ability through the process.

An emissions control device 21 may be used to remove acid gases from theflue gas stream, such as SO2 and HCL, resultant from the combustioncycle. This emissions control device may be in the form of a dryinjection system or a dry spray absorption system. In an embodiment,these systems may consist of multiple sections or devices: a device toblend the dry sorbent media with liquid in order to develop a sorbentslurry or a device to directly incorporate dry sorbent, a device tometer and inject said sorbent slurry or dry sorbent into the flue gasstream, a device to control the flow rate of the flue gas streamenabling efficient absorption of acid gases into the sorbent, acollection vessel for removal of precipitated dried reacted products(acid gases and sorbent slurry), or a secondary system (bag house) tocapture absorbed dry reacted product or residuals that remain within theflue gas stream (downstream of the dry spray absorber or dry injectionsystem).

A bag house 22 may be provided for filtering remaining particulate fromthe flue gas, and may comprise a pulse-jet type bag house. Flue gasesentering the bag house are cleaned by filtering the particulate throughthe bags in the bag house. A high-pressure blast of air is used toremove dust from the bag. Due to its rapid release, the blast of airdoes not interfere with contaminated flue gas flow. Therefore, thepulse-jet bag house can operate continuously with high effectiveness incleaning the exiting flue gas.

A clean flue gas stack 23 provides the normal exit point from theprocess and one of many emission testing points that may be utilized tomaintain constant regulatory compliance.

A SCADA system can be utilized to ensure full process control to allaspects of the biogasification system. As such, typical thermocouples,pressure instruments, level switches, gas analysis, moisture metering,flow meters and process instrumentation are used to provide ‘real time’feedback, and automated adjustment of actuators to ensure optimum andefficient operation.

FIG. 1B shows an alternate embodiment of the overall flow diagram ofFIG. 1A. In this embodiment, two heat exchangers are utilized ratherthan three as in the embodiment of FIG. 1A. In accordance with thisembodiment, the second heat exchanger 17 can be utilized for additionalenergy recovery streams (typically wasted) and return the energy to theprocess for optimum performance efficiency. The second heat exchanger 17may be used as a biogasifier oxidant air and gas pre-heater. The sourceof the oxygen for the disclosed gasification system may be designed tocome from air or flue gas, or mixtures thereof. In an embodiment, thedisclosed system may be configured such that the biogasifier oxidant airand gas, independently or in a mixed configuration, may be pre-heated asa means of energy recovery to further optimize the gasification systemperformance efficiency.

FIG. 2 shows a side view illustrating a fluidized bed biogasifier 200 inaccordance with an embodiment of the invention. In an embodiment, thefluidized bed biogasifier 200 includes a bubbling type fluidized bed.The fluidized bed may utilize sand or other fluidizing media. A bubblingreactor bed section 204 is filled with media. Fluidized fuel feed inlets201 in the reactor bed section 204 receive sludge at 40°-250° F., e.g.,215° F., and a flue gas inlet 203 in the bubbling bed receives flue gasat, e.g., 600° F. The flue gas can be fed to the bubbling bed via a gasdistributor (FIG. 3). An oxygen monitor 209 may be provided incommunication with the flue gas inlet 203 to monitor oxygenconcentration in connection with controlling oxygen levels in thegasification process. An inclined or over-fire natural gas burner (notvisible) located on the side of the reactor vessel receives a naturalgas and air mixture via a port 202 at, e.g., 77F and uses the same asfuel. View ports 206 and a media fill port 212 are provided.

A freeboard section 205 may be provided between the fluidized bed in thereactor bed section 204 and the outlet 210 of the biogasifier. As thebiosolids thermally decompose in the fluidized bed and then rise throughthe reactor vessel, they are transformed through a reactor bed section.The fluidizing medium in the fluidized bed is disentrained from theproducer gas in a disengaging zone. A cyclone separator 207 may beprovided to separate material exhausted from the fluidized bed reactorinto clean producer gas for recovery and ash for disposal. An optionalash grate, discussed below, may be fitted below the biogasifier vesselfor bottom ash removal. A producer gas control 208 monitors oxygen andcarbon monoxide levels in the producer gas and controls the processaccordingly.

A number of thermocouple probes are placed in the biogasifier to monitorthe temperature profile throughout the biogasifier. Some of the thermalprobes are placed in the fluidized bed section 204 of the biogasifierwhile others are placed in the freeboard section 205 of the biogasifier.The thermal probes placed in the fluidized bed are used not only tomonitor the bed temperature, but are also control points that arecoupled to the biogasifier air system in order to maintain a certaintemperature profile in the bed of fluidizing media. There are also anumber of additional instruments placed in the biogasifier to monitorthe pressure differential across the bed and the operating pressure ofthe biogasifier in the freeboard section. These additional instrumentsare used to monitor the conditions within the biogasifier as well to ascontrol other ancillary equipment and processes to maintain the desiredoperating conditions within the biogasifier. Examples of such ancillaryequipment and processes include, e.g., the cyclone, thermal oxidizer andrecirculating flue gas and air delivery processes.

FIG. 3 shows a perspective view illustrating the gas distributor 302 ofthe biogasifier in accordance with an embodiment of the invention. Aflue gas and air inlet 203 feeds flue gas and air to an array of nozzles301. Each of the nozzles includes downwardly directed ports in cap 303such that gas exiting the nozzle is initially directed downward beforebeing forced upward into the fluidized bed in the reactor bed section204 (FIG. 2). An optional ash grate under the gasifier may be used as a‘sifting’ device to remove any agglomerated particles so that thefluidizing media and unreacted char can be reintroduced into thebiogasifier for continued utilization.

With reference to FIGS. 1, 2 and 4A-4B, startup and operation of thebiogasifer will now be described. This example is based on a feed rateof 1850 lbs/hr of biosolids at 90% solids.

Start-Up of the Biogasifier

The inline or over-fire natural gas burner is ignited and a slip streamof gasifier air mixed with natural gas is combusted and directed ontothe top of the fluidized bed to heat up the refractory lined biogasifierreactor.

The bed begins to warm up and fluidization becomes more aggressive asthe fluidizing media and fluidizing gas heat up.

The hot gas from the natural gas burner then passes into the uppersection of the biogasifier and into the cyclone. This begins the warm upof the refractory in these components.

The induced draft fan is turned on and the hot gas is drawn into theoxidizer, through the heat exchangers, into the scrubbing system, andthen up the stack. This continues the warm up process of the plant.

This process is continued until the biogasifier bed temperature is above900° F. and the refractory temperature in the cyclone and oxidizer areat least 600° F.

The flue gas recycle blower is then started slowly and additionalfluidizing gas is introduced into the biogasifier. This increases thefluidization regime and increases the effectiveness of the cyclone atremoving particulate.

The dried sludge at 40°-250° F. from the dryer is now slowly added tobegin heat generation within the bed of fluidizing media and thetemperature is monitored closely at this point, operating within atemperature range of 1150°-1600° F. This reduces the possibility ofagglomerating or “clinkering” of the ash.

The biogasification reaction begins and producer gas is produced, whichentrains some solids (ash and unreacted carbon) into the cyclone whereparticulate matter larger than 10 microns is removed by the cyclone.

The freeboard or disengaging zone allows the largest sized particles todecelerate and fall back to the bed of the biogasifier.

The solids separated by the cyclone are emptied from the base of thecyclone into a specially designed separator where the fine ash (and somecarbon) is taken out as waste and the remainder is recycled back to thebiogasifier. This increases the overall conversion of fuel to close to100%.

As the temperature increases and approaches the target operatingtemperature, preferably in the range of 1150°-1600° F., the gasificationand pyrolysis reactions are at the desired levels.

Operation of the Biogasifier

The oxygen starved environment in the biogasification process may beachieved by controlling the level of oxidant air and gas entering intothe reactor. Oxidant air, or ambient air, consists of approximately23.2% oxygen by weight. Oxidant gas, or recycled flue gas, consists ofoxygen levels that may vary from 5% to 15% by weight.

The temperature within the biogasifier may be controlled by adjustingthe amount of oxidant air and gas entering into the reactor.

The ratio of oxygen entering the biogasifier supplied by oxidant air andgas in relation to the stoichiometric oxygen required for completecombustion of the biosolids fuel is referred to as the equivalence ratioand may be used as a means of regulating temperature within thebiogasification unit. Gasification occurs by operating a thermochemicalconversion process with an oxygen-to-fuel equivalence ratio between 0.1and 0.5 relative to complete stoichiometric combustion.

Depending on the selected biosolids fuel feed rate the blend of Oxidantair to Oxidant gas will vary in order to sustain a total mass levelwithin the biogasifier required to operate the cyclone within anacceptable range of efficiency while meeting the targeted equivalenceratio selected in order to control the process temperature.

Bed pressure within the biogasifier reaction zone may be monitored as ameans of controlling the pressure of the biogasifier oxidant air and gasentering the biogasification unit and thereby ensuring continuousfluidization of the reactor bed media.

In the biogasifier 200, the oxygen is used to control the temperature,extend the reaction and overall generation of the producer gas, andtypically these reactions are:C+O2→CO2 Exothermic2H+0.5O2→H2O Exothermic

The heat generated heats up the incoming biosolids feed, biogasifier airand recycled flue gas streams to the reaction temperature.

It also provides the heat to complete the mostly endothermicbiogasification reactions that take place in the biogasifier, namely:C+H2O→CO+H2 EndothermicCO2+C→2CO EndothermicH2O+CO→H2+CO2 ExothermicC+2H2→CH4 Exothermic

If excessive oxygen is introduced, the quality of the producer gas willbe reduced. Maintaining a uniform operating temperature across thefluidized bed of the biogasifier is key to operational success.

Based on the biogasifier air and flue gas streams, and fuel propertiesplus conversion extent, the superficial velocity of the gas in the bedand freeboard section of the biogasifier can be calculated. Thisdetermines what size particles of fluidizing media (e.g., sand), ash,and biosolids are entrained out with the producer gas. Many factors areused to calculate superficial velocity inside the reactor, including airflow rate, recycled flue gas flow rate, reactor physical geometry suchas cross sectional area and shape, sewage sludge fuel composition, fuelfeed rate, and fuel conversion extent.

Using recycled flue gas, which has a lower percentage of oxygen thanair, allows not only a better way to control the oxygen and hencetemperature in the biogasifier, but also allows greater control of thesuperficial velocity of the producer gas in both the bed and freeboardsections of the biogasifier.

The biosolids, even though they are 90% solids, include clusteredparticles that are partially reacted and can be carried out of thebiogasifier prematurely. The usual once through fuel conversion is about85%, so 15% unreacted carbon may be carried into the cyclone.

The particle density of the inert ash is about 2160 kg/m³, thefluidizing media is around 2600 kg/m³, and the sludge particles around600 kg/m³.

By choosing a superficial gas velocity of 1.1 m/s in the freeboardsection, for example, it is possible to entrain out ash particles of 200microns and fluidizing media particles of 100 microns in the fluidizedbed. The target particle size range of the fluidizing media is withinthe range of 400-900 microns, so that none is lost to the cycloneseparator 207.

About 50% of the ash has particles that are less than 200 microns andwill be carried over to the cyclone separator 207, with about 50% beingleft in the bed.

Some unreacted carbon is carried into the cyclone separator 207 withparticle sizes ranging from 10 to 300 microns. When the solids areremoved from the bottom of the cyclone, the ash and unreacted carbon canbe separated and much of the unreacted carbon recycled back into thebiogasifier, thus increasing the overall fuel conversion to at least95%. Ash accumulation in the bed of fluidizing media may be alleviatedthrough adjusting the superficial velocity of the gases rising insidethe reactor. Alternatively, bed media and ash could be slowly drainedout of the gasifier base and screened over a grate before beingreintroduced back into the biogasifier. This process can be used toremove small agglomerated particles should they form in the bed offluidizing media and can also be used to control the ash-to-media ratiowithin the fluidized bed.

Operation of the Thermal Oxidizer

When the dried biosolids feed to the biogasifier is started, theoxidizer air is turned on and passed into the heat exchanger 19 (FIG. 1)where it is slowly preheated before being added to the oxidizer 12. Theoxidizer may be fitted with an ignition source and when the producer gasreaches the reaction chamber, the oxidation commences and significantcombustion heat is produced.

During normal operation, the temperature of the flue gas leaving theoxidizer is, for example between 1800° and 2200° F. The sensible heatfrom this stream is transferred via the pre-dryer and dryer heatexchangers and the flue gas temperature is reduced from 1800° F. toabout 550° F.

When required, the heat exchanger 17 is used to pre-heat the biogasifiergas stream prior to injection into the biogasifier.

The flue gas is now about 500° F. and between 20 and 30% of the flue gasis recycled back to the biogasifier using the flue gas recycle blower.The remainder passes into heat exchanger 3 where the oxidized air may bepre-heated to at least 300° F. When the oxidizer air is pre-heated to300° F. and over, less energy from combustion of producer gas may beused to heat the air up to the temperature that the thermal oxidizeroperates at, thus increasing the performance of the oxidizer.

Using the oxygen level defined at the oxidizer inlet, the oxygen levelin the flue gas can be adjusted by increasing/decreasing the amount ofair being fed to the oxidizer. Oxygen levels can be controlled bylimiting the use of external or ambient air and increasing an amount ofrecirculated flue gas in order to maintain a required temperatureprofile at the exit of the oxidizer.

This has the effect of increasing/decreasing the excess air being fed tothe oxidizer but is never below the minimum level of excess air required(>25%), to ensure complete combustion and low emissions.

Biogasifier Reactor Sizing

The following provides a non-limiting example illustrating computationof the best dimensions for a bubbling fluidized bed biogasificationreactor in accordance with an embodiment of the invention. Thebiogasifier, in this example, is sized to accommodate two specificoperating conditions:

-   -   1. The current maximum dried biosolids output generated from the        dryer with respect to the average solids content of the        dewatered sludge supplied to the dryer from the existing belt        press dewatering unit, and    -   2. The future maximum dried biosolids feed rate that the dryer        will have to deliver to the biogasifier if the overall biosolids        processing system has to operate without consumption of external        energy, e.g., natural gas, during steady state operation with        25% solids content dewatered sludge being dried and 5400 lb/hr        of water being evaporated from the sludge.

The first operating condition corresponds to the maximum output of driedsewage sludge from the dryer if, e.g., 16% solids content sludge isentering the dryer, and 5400 lb/hr of water is evaporating off thesludge. This corresponds to a biosolids feed rate of 1168 lb/hr ofthermally dried biosolids at 10% moisture content entering the gasifier.The second operating condition corresponds to the maximum amount ofdried biosolids (dried to 10% moisture content) that the drier canproduce if 25% solids content dewatered biosolids is fed into the drier.A solids content of 25% represents the estimated extent of dewateringthat is required to make the drying load equal to the amount of thermalenergy which can be recovered from the flue gas and used to operate thedryer. If biosolids below 25% solids content are processed in the dryer,an external heat source can be expected to be required to supplement thedrying process. The second condition corresponds to the gasifier needingto process 2000 lb/hr of 10% moisture content biosolids.

FIG. 5 shows a non-limiting example of the biogasifier internaldimensions in accordance with the invention in an embodiment. Thedimensions shown satisfy the operational conditions that are outlinedbelow. Tables 1-3 below show a non-limiting example of physicalparameters taken into consideration when setting the biogasifierdimensions. Tables 1 and 2 show fuel composition results in accordancewith an ultimate analysis and a proximate analysis, respectively.

TABLE 1 Element wt %_(drybasis) C: 41.9 H: 5.5 O: 21.6 N: 6.5 S: 1.0 Ash(assume 100% Si): 23.5 100

TABLE 2 wt %_(drybasis) Volatile matter: 65.4 Fixed carbon: 11.1 Ash23.5

TABLE 3 Case 1: Low Case 2: High Biosolids Biosolids Feed Rate of FeedRate of Parameter: 1168 lb/hr 2000 lb/hr Average temperature 1450° F. or1450° F. or in gasifier 788° C. 788° C. (media bed & freeboard) Moisturecontent of the  10%  10% biosolids into the gasifier Calorific value ofthe 7600 Btu/lb 7600 Btu/lb biosolids (HHV dry basis) Estimated averagefuel 300 μm 300 μm particle size Estimated average ash 200 μm 200 μmparticle size Target equivalence ratio 0.3 0.3 Amount of gasification1825 lb/hr 3125 lb/hr air required Volume flow rate of 0.17 m³/s 0.29m³/s gasification air (at S.T.P) Assumed volatile matter 100% 100%conversion to producer gas Assumed fixed carbon  50%  50% conversion toproducer gas Producer gas mass flow rate 0.339 kg/s or 0.580 kg/s or2630 lb/hr 4592 lb/hr Producer gas volume 1.02 m³/s at 1.74 m³/s at flowrate (actual) 1450° F. 1450° F. Producer gas volume 2352 ACFM at 4029ACFM at flow rate (actual) 1450° F. 1450° F. Producer gas volume 0.29m³/s 0.49 m³/s flow rate (S.T.P) Producer gas volume 661 SCFM 1133 SCFMflow rate (S.T.P) Estimated fuel 600 kg/m³ 600 kg/m³ particle densityEstimated ash 2160 kg/m³ 2160 kg/m³ particle density Estimated media(sand) 2600 kg/m³ 2600 kg/m³ particle density Selected Ø I.D of bedsection 3.75 ft or 3ft, 9″ 3.75 ft or 3ft, 9″ Media bed depth 3 ft 3 ft(unfluidized) Height of reactor bed-section 4.5 ft 4.5 ft Superficialvelocity of gas 0.99 m/s 1.70 m/s above media bed surface Superficialvelocity from 0.16 m/s 0.28 m/s gasification air at S.T.P Minimumfluidization 0.16 m/s 0.16 m/s velocity, U_(mf) based on air at S.T.PAverage media (sand) ~700 μm ~700 μm particle size in the bed Pressuredrop across 1.7 psi 1.7 psi incipiently fluidized bed (gas distributorto bed surface). Selected Ø I.D of 4.75 ft or 4 ft, 9″ 4.75 ft or 4ft,9″ freeboard section Superficial velocity of 0.62 m/s 1.06 m/s gas infreeboard T.D.H for particles ~11.9 ft ~21.8 ft Maximum size of fuelparticle 195 μm 390 μm which will get entrained Maximum size of ashparticle 100 μm 200 μm which will get entrained Maximum size of 85 μm180 μm media particle which will get entrained Energy input to gasifier2.34 MW_(th) 4.00 MW_(th) (HHV basis) Equivalent grate energy 2.28MW_(th)/m² 3.90 MW_(th)/ release rate m² Equivalent grate energy 722,596Btu_(th)/ft² 1,237,322 Bt_(uth)/ft² release rate (~1 million Btu per ft²of bed area) Reactor diameter to fuel 1.61 ft/[500 0.94 ft/[500 feedrate ratio (~1ft diameter lb/hr fuel] lb/hr fuel] for every 500 lb/hrfuel)

With continued reference to FIG. 5, one factor in determiningbiogasifier sizing is the bed section internal diameter. The role of thebed section of the reactor is to contain the fluidized media bed. Thedriving factor for selecting the internal diameter of the bed section ofthe gasifier is the superficial velocity range of gases, which varieswith different reactor internal diameters. The internal diameter has tobe small enough to ensure that the media bed is able to be fluidizedadequately for the given air, recirculated flue gas and fuel feed ratesat different operating temperatures, but not so small as to create suchhigh velocities that a slugging regime occurs and media is projected upthe freeboard section. The media particle size can be adjusted duringcommissioning to fine tune the fluidizing behavior of the bed. In thepresent, non-limiting example, an average media (sand) particle size of700 μm was selected due to its ability to be fluidized readily, but alsoits difficulty to entrain out of the reactor. The most difficult time tofluidize the bed is on start up when the bed media and incoming gasesare cold. This minimum flow rate requirement is represented by theminimum fluidization velocity (U_(mf)) values displayed in the previoustable.

Another factor in determining biogasifier sizing is the freeboardsection internal diameter. The freeboard region of the biogasifierallows for particles to drop out under the force of gravity. Thediameter of the freeboard is selected with respect to the superficialvelocity of the gas mixture that is created from different operatingtemperatures and fuel feed rates. The gas superficial velocity must begreat enough to entrain the small ash particles, but not so great thatthe media particles are entrained in the gas stream. The extent of freshfuel entrainment should also be minimized from correct freeboard sectionsizing. This is a phenomenon to carefully consider in the case ofbiosolids gasification where the fuel typically has a very fine particlesize. Introducing the fuel into the side of the fluidized bed below thefluidizing media's surface is one method to minimize fresh fuelentrainment. This is based on the principle that the fuel has to migrateup to the bed's surface before it can be entrained out of thebiogasifier, and this provides time for the gasification reactions tooccur. In the present non-limiting example, a reactor freeboard diameterof 4 ft, 9″ is chosen in an effort to maintain gas superficialvelocities high enough to entrain out ash, but prevent entrainment ofsand (or other fluidizing media) particles in the bed.

A further factor in determining biogasifier sizing is the media beddepth and bed section height. In general, the higher the ratio of mediato fuel in the bed, the more isothermic the bed temperatures are likelyto be. Typically, fluidized beds have a fuel-to-media mass ratio ofabout 1-3%. The amount of electrical energy consumed to fluidize themedia bed typically imparts a practical limit on the desirable depth ofthe media. Deeper beds have a higher gas pressure drop across them andmore energy is consumed by the blower to overcome this resistance to gasflow. A fluidizing media depth of 3 ft is chosen in this example basedon balancing the blower energy consumption against having enough mediain the bed to maintain isothermal temperature and good heat transferrates. The height of the bed section of the reactor in this non-limitingexample is based on a common length-to-diameter aspect ratio of 1.5,relative to the depth of the fluidizing media.

Another factor in determining biogasifier sizing is the height of thefreeboard section. The freeboard section is designed to drop outparticles under the force of gravity. As one moves up in elevation fromthe bed's surface, the particle density decreases, until at a certainelevation, a level known as the Transport Disengaging Height (TDH) isreached. Above the TDH, the particle density entrained up the reactor isconstant. Extending the reactor above the TDH adds no further benefit toparticle removal. For practical purposes loft is selected in thisnon-limiting example.

While the invention has been particularly shown and described withreference to a preferred embodiment thereof, it will be understood bythose skilled in the art that various changes in form and details may bemade therein without departing from the spirit and scope of theinvention.

What is claimed is:
 1. A method for gasifying biosolids obtained fromsewage sludge, comprising: receiving biosolids in a fluidized bedbiogasifier, the biogasifier having: a reactor vessel; a freeboardsection having a diameter of at least 57 inches and a height of at least10 feet; a feeder for feeding biosolids into said reactor vessel, saidfeeder being configured to feed said biosolids into said reactor vesselat a biosolids fuel feed rate during steady-state operation of thebiogasifier; and a fluidized bed in a bed section of said reactorvessel, said fluidized bed having a diameter of at least 45 inches;wherein the freeboard section has a greater diameter than the fluidizedbed, wherein the ratio of the freeboard section diameter to thefluidized bed diameter is at least 57:45; and further wherein a ratio ofa height of the bed section of the reactor vessel to a depth of thefluidized bed is 1.5; introducing gas to said fluidized bed reactor; andheating and reacting said biosolids inside said biogasifier, wherebybiosolids are gasified.
 2. The method for gasifying biosolids inaccordance with claim 1, wherein a superficial velocity of producer gasinside the freeboard section during steady state operation is between0.1 m/s (0.33 ft/s) and 3 m/s (9.84 ft/s).
 3. The method for gasifyingbiosolids in accordance with claim 1, wherein a thermochemicalconversion process occurs within the biogasifier at a temperaturebetween 900° F. (482.2° C.) and 1700° F. (926.7° C.).
 4. The method forgasifying biosolids in accordance with claim 1, wherein said fluidizedbed reactor comprises a bubbling fluidized bed reactor.
 5. The methodfor gasifying biosolids in accordance with claim 3, wherein saidtemperature range is between 1150° F. (621.1° C.) and 1600° F. (871.1°C.).
 6. The method for gasifying biosolids in accordance with claim 1,wherein a superficial velocity of the producer gas inside the freeboardsection during steady state operation is between 0.5 m/s and 2.5 m/s. 7.The method for gasifying biosolids in accordance with claim 1, furthercomprising the step of using a cyclone separator to separate materialexhausted from said fluidized bed reactor into producer gas and ash. 8.The method for gasifying biosolids in accordance with claim 3, whereinthe thermochemical conversion process occurs with an oxygen-to-fuelequivalence ratio between 0.1 and 0.5.
 9. The method for gasifyingbiosolids in accordance with claim 3, wherein the thermochemicalconversion process occurs with an oxygen-to-fuel equivalence ratiobetween 0.15 and 0.35.
 10. A method for gasifying biosolids obtainedfrom sewage sludge, comprising: receiving biosolids in a biogasifierapparatus, the apparatus having: a fluidized bed biogasifier forgasifying biosolids to produce a producer gas; a staged-combustionthermal oxidizer for thermally oxidizing said producer gas; and arecycle flue gas loop connecting the staged-combustion thermal oxidizerto the fluidized bed biogasifier; wherein the fluidized bed gasifiercomprises: a reactor vessel; a freeboard section having a diameter of atleast 57 inches and a height of at least 10 feet; a feeder for feedingbiosolids into said reactor vessel, said feeder being configured to feedsaid biosolids into said reactor vessel at a biosolids fuel feed rateduring steady-state operation of the biogasifier; and a fluidized bed ina bed section of said reactor vessel, said fluidized bed having adiameter of at least 45 inches; wherein the freeboard section has agreater diameter than the fluidized bed; wherein the ratio of thefreeboard section diameter to the fluidized bed diameter is at least57:45; and further wherein the recycle flue gas loop is configured torecycle flue gas to the fluidized bed biogasifier; introducing gas tosaid fluidized bed reactor; and heating and reacting said biosolidsinside said biogasifier, whereby biosolids are gasified.
 11. The methodfor gasifying biosolids in accordance with claim 10, wherein asuperficial velocity of producer gas inside the freeboard section duringsteady state operation is between 0.1 m/s (0.33 ft/s) and 3 m/s (9.84ft/s).
 12. The method for gasifying biosolids in accordance with claim10, wherein a thermochemical conversion process occurs within thebiogasifier at a temperature between 900° F. (482.2° C.) and 1700° F.(926.7° C.).
 13. The method for gasifying biosolids in accordance withclaim 10, wherein said fluidized bed reactor comprises a bubblingfluidized bed reactor.
 14. The method for gasifying biosolids inaccordance with claim 12, wherein said temperature range is between1150° F. (621.1° C.) and 1600° F. (871.1° C.).
 15. The method forgasifying biosolids in accordance with claim 10, wherein a superficialvelocity of the producer gas inside the freeboard section during steadystate operation is between 0.5 m/s and 2.5 m/s.
 16. The method forgasifying biosolids in accordance with claim 10, further comprising thestep of using a cyclone separator to separate material exhausted fromsaid fluidized bed reactor into producer gas and ash.
 17. The method forgasifying biosolids in accordance with claim 12, wherein thethermochemical conversion process occurs with an oxygen-to-fuelequivalence ratio between 0.1 and 0.5.
 18. The method for gasifyingbiosolids in accordance with claim 12, wherein the thermochemicalconversion process occurs with an oxygen-to-fuel equivalence ratiobetween 0.15 and 0.35.